Hydroconversion of petroleum oils

ABSTRACT

In a dual catalyst system such as a petroleum oil hydrotreating - hydrocracking combination in which the first reaction zone is used to control the amount of catalyst deactivant passing into the second reaction zone, and the second reaction zone is a hydroconversion zone the temperature differential between the end-of-run temperature of each catalyst is maintained throughout the on-stream period and the reaction temperatures are regulated to obtain the desired amount of conversion.

United States Patent 91 Hahn [54] HYDROCONVERSION OF PETROLEUM OILS [75]Inventor: Frederick K. Hahn, Port Arthur,

Tex.

[73] Assignee: Texaco, Inc., New York, NY.

[22] Filed: Aug. 16, 1971 [2l] Appl. No.: 172,259

[52] US. Cl ..208/89, 208/59 [51] Int. Cl. ..C 10g 23/00 [58] Field ofSearch ..208/89, 59

[56] References Cited UNITED STATES PATENTS 3,338,819 8/1967 Wood..208/89 3,506,568 4/l970 Annesser et al. ..208/89 1 Mar. 27, 19733,598,719 8/1971 White ..208 59 Primary ExaminerDelbert E. GantzAssistant Examiner-James W. Hellwege Att0rney'lhomas H. Whaley et al.

[57] ABSTRACT In a dual catalyst system such as a petroleum oilhydrotreating hydrocracking combination in which the first reaction zoneis used to control the amount of catalyst deactivant passing into thesecond reaction zone, and the second reaction zone is a hydroconversionzone the temperature differential between the end-of-run temperature ofeach catalyst is maintained throughout the on-stream. period and thereaction temperatures are regulated to obtain the desired amount ofconversion.

13 Claims, No Drawings HYDROCONVERSION OF PETROLEUM OILS This inventionrelates to the conversion of hydrocarbons. More particularly, it isconcerned with the catalytic conversion of hydrocarbons using twocatalysts in series in which the first catalyst is directed to reducingthe amount of materials in the feed which are deleterious to theactivity of the second catalyst or to converting them into materialswhich have less or no undesirable effects. In a specific embodiment itis directed to the conversion of heavier petroleum hydrocarbons such asgas oils into lighter petroleum hydrocarbons such as motor and jet fuelsusing a two catalyst system in which materials which are deleterious tothe activity of the second catalyst are removed or converted by thefirst catalyst to less deleterious materials.

It is customary in petroleum refining to carry out various hydrogenationreactions in the presence of a catalyst. However, frequently there arematerials present in the charge stock which have a harmful effect on thecatalyst. It is therefore necessary, for prolonged operation, either toremove these materials or to convert them to materials which have lessor no undesirable effects. For example, in the hydrocracking ofpetroleum oils the catalyst activity is impaired by the presence oforganic nitrogen compounds in the feedstock. However, hydrocrackingcatalysts containing zeolitic compounds in the support, althoughsomewhat susceptible to deactivation by organic nitrogen compounds, arequite resistant to deactivation due to the presence of ammonia in thefeed. Although ammonia tends to suppress the activity of zeolitehydrocracking catalysts somewhat, the effect is reversible simply byremoving the ammonia, and does not increase with time. Consequently, ithas become conventional to pretreat the feedstock to a hydrocrackingunit by contacting it in the presence of hydrogen with a hydrotreatingcatalyst under conditions to convert the organic nitrogen compounds toammonia and to pass the entire effluent from the hydrotreating zone tothe hydrocracking zone. In such a procedure, the feed, usually virgingas oil, coker gas oil, or catalytic cycle gas oil, is contacted atelevated temperature and pressure with hydrogen in the presence of ahydrogenation catalyst to convert the catalyst deactivant such asorganic nitrogen or polycyclic aromatic compounds into ammonia ormonocyclic aromatic compounds respectively. The effluent from thehydrotreating unit is then introduced into the hydrocracking unit at thedesired temperature and pressure and the pretreated oil is subjected tohydrocracking therein.

Ordinarily the hydrotreating unit is operated at conditions to produce aproduct having a specified organic nitrogen content. For example, if itis desired to maintain the organic nitrogen content of the petroleumfeed to the hydrocracking unit in the range of 40-50 ppm organicnitrogen, then the effluent from the hydrotreating unit is sampledperiodically and analyzed to determine its organic nitrogen content. Thereaction conditions in the hydrotreating unit are then adjusted tomaintain the desired level. As the on stream period progresses, thecatalyst activity diminishes usually due to the presence of carbondeposited thereon and to compensate for the decreased activity itbecomes necessary to increase the severity of the reaction conditionsgenerally by increasing the reaction temperature.

Eventually, however, the catalyst bed is raised to a temperature abovewhich it is no longer advantageous to operate and the unit is shut downfor catalyst regeneration. This is done by purging the reactor with aninert gas to remove combustible materials such as residual hydrocarbonand hydrogen. The regeneration of the catalyst is then effected byintroducing a gas containing a small amount of oxygen and the carbon isgradually burned off the catalyst under carefully controlled ternperature conditions until the effluent gas shows that there has beenessentially complete removal of carbon. The introduction of theoxygen-containing gas is then terminated, the reactor purged with aninert gas, hydrogen is introduced into the reactor, the catalyst isresulfided if desired and then brought to reaction temperature. The onstream period is then resumed by introduction of the hydrocarbon oilfeed.

Similarly, in the operation of the hydrocracking unit, the operatingconditions such as temperature and pressure are set to obtain thedesired amount of conversion of the charge stock to lighter boilingmaterials. As in the case of the hydrotreating unit, the catalystgradually loses its activity due to, among other things,

the deposition of carbon thereon and to compensate for the decreasedactivity and to maintain the desired amount of conversion, the catalysttemperature is gradually raised. Eventually, a temperature is reachedabove which it is no longer desirable to operate and the hydrocrackingunit is shut down for regeneration in a manner, similar to thehydrotreating unit. Unfortunately, the operating ranges and deactivationrates of the catalysts are different and seldom, if ever, does it occurthat both the hydrotreating and hydrocracking catalysts requireregeneration at the same time. Usually, it is necessary to shut down theentire process while the catalyst in one unit is on regeneration andsubsequently to shut down the entire process while the catalyst in theother unit is being regenerated. This means that frequently although oneunit is still capable of being operated, the entire process is shut downfor the regeneration of the catalyst in the other unit. This is anuneconomical method of operation.

It is, therefore, an object of the present invention to operate atwo-catalyst hydrocarbon conversion process more efficiently thanheretofore. Another object is to operate a two-stage catalyst conversionprocess so that the catalysts reach their end-of-run temperaturesubstantially simultaneously. Another object is to maintainsubstantially the same average temperature differential between the twocatalyst beds. Still another object is to reduce the non-productivedowntime of a two-catalyst hydroconversion process. These and otherobjects will be obvious to those skilled in the art from the followingdisclosure.

According to our invention, in a hydrocarbon conversion process which isconducted using two catalysts in series and in which the first catalystis a pretreating catalyst for controlling the amount of catalystdeactivant entering the second stage, and the second catalyst is aconversion catalyst, the operation is improved by maintainingsubstantially constant a predetermined average catalyst bed temperaturedifferential between the first andsecond catalyst beds and controllingthe catalyst bed temperatures by the amount of conversion effected inthe second stage.

Although the process is applicable to several types of hydrocarbonconversion processes, it is best described in a specific embodiment, asfor example, the hydrocracking of gas oils to motor and jet fuels.

To establish the operation of the process of our invention, theend-of-run temperature for each catalyst is determined and thetemperature differential between these end-of-run temperatures ismaintained substantially constant throughout the on stream period. Forexample, if the end-of-run temperature for the second or hydrocrackingstage catalyst is 745 F. and for the hydrotreating catalyst is 775 F.then a temperature differential of approximately 30 F. is maintainedbetween the catalyst beds throughout the on stream period. Ordinarilythe end-of-run temperatures are selected basis the changes in productdistribution and/or product quality that occur as the reactiontemperature is increased to compensate for catalyst aging. Generally, asthe temperature is increased the product distribution will becomelighter, i.e. larger amounts of light hydrocarbons such as methane,ethane, and propane will be produced while the yields of jet fuel andheavy naphtha will be decreased. Also, the aromatic concentration in theproducts will generally increase as temperature is increased. Theend-of-run temperature is then chosen as the temperature above which theproduct distribution is determined to be unacceptable, or possibly whenhydrocracking to produce jet fuels or diesel fuels, the temperatureabove which the aromatic. content of the product is unacceptable. Thistemperature can be determined experimentally. Occasionally, particularlyin commercial units, the end-of-run temperature may be set basis designconsideration, e.g., the temperature at which heater and/or heatexchanger fouling becomes unacceptable.

in the practice of our invention, if two reactors are used the on-streamperiod is inaugurated by introducing the hydrocarbon charge withhydrogen into the first reactor and passing the effluent from the firstreactor through the second reactor at the predetermined conditions ofpressure, space velocity and hydrogen rates. The catalysts may bepreheated to substantially operating conditions prior to theintroduction of the reactant stream. The catalyst bed temperatures aregradually increased while maintaining the proper temperaturedifferential therebetween until the desired amount of conversion isobtained in the second stage. in the case of a gas oil feed, this may befor example a conversion of 50-100 percent to jet fuel and lighterboiling material. The reaction conditions are then stabilized. Thetemperature in the second stage may be higher or lower than wouldordinarily be used in the prior art hydrocracking processes to process afeed having the same content of catalyst deactivant. For example, if inthe prior art processes, it has been customary to maintain the organicnitrogen feed of the hydrocarbon charge to the second zone at between 40and 50 ppm nitrogen, the starting temperature of the first stage mayactually produce a feed to the second zone having a nitrogen content ofbetween and ppm. However, this will be compensated for by maintainingthe second conversion zone temperature at below what would ordinarily bethe start of run temperature for the particular catalyst and feed stock.Since the feed to the hydrocracking unit is lower in catalystdeactivant, the

same amount of conversion can be obtained at lower temperature than witha feed containing, for example, between 40 and 50 ppm nitrogen. Thus,once the processing conditions have been established, the entire processis governed by the temperature of the second catalyst bed necessary toobtain and maintain the desired amount of conversion.

As the conversion drops off because of loss of activity of the catalyst,the temperature in both reactors in increased. This increases both theconversion in the second reactor and also the amount of catalyst poisonremoval in the first reactor. if the first reactor is being operated attoo low a temperature thereby permitting an unduly high amount ofnitrogen to be introduced with the feed to the second stage this willaffect the activity of the second stage catalyst and reduce the amountof conversion. By raising the temperature in both reactors slightly, theconversion of the organic nitrogen to ammonia in the first reactor willbe increased, thereby reducing the amount of organic nitrogen going tothe second reactor which in turn results in increased conversion. If theconversion increases beyond the desired level, a cutback in thetemperature in both the first and second reactors will permit moredeactivant to pass from the first reactor to the second reactor whichcoupled with the lower temperature in the second reactor will result inreduced activity of the second catalyst with a return of the conversionin the second reactor to the desired level.

Although in the preceding paragraphs, the process has been described asbeing carried out with each bed of catalyst situated in its own reactorvessel, it is possible to have both catalyst beds in the same reactorvessel. Whether the system contains only one reactor vessel or tworeactor vessels, the greatest advantage is obtained if the entireeffluent from the first catalyst bed is sent to the second catalyst bedalthough in some instances particularly where the charge to the firstcatalyst has an extremely high deactivant content it may be preferred toremove a portion of the reactant gas stream between beds. Preferably thereactant stream is passed downflow through fixed beds of particulatecatalyst. However, it is possible to have countercurrent hydrogen-oilflow or to flow the reactants upwardly through a fixed or fluidized bedof catalyst.

By following our procedure, it is no longer necessary to sample theeffluent from the hydrotreating unit periodically and to determineanalytically the nitrogen content of the hydrocarbon feed to thehydrocracking unit. The process conditions are controlled simply by theamount of conversion effected in the second stage which can easily bedetermined by a method as simple as reading a gauge.

The process of this invention is applicable to various petroleumfractions such as naphthas, kerosenes, virgin gas oils, cycle gas oils,vacuum gas oils, residua, shale oil, tar sand oil and the like.

The hydrogen used in the process of our invention need not necessarilybe pure. The hydrogen content of the hydrogenating gas should be atleast about 60 percent and preferably is at least about percent byvolume. Suitable sources of hydrogen are catalytic reformer by-producthydrogen and hydrogen produced either by the partial combustion ofhydrocarbonaceous material or steam reforming of light hydrocarbonsfollowed by shift conversion and CO removal. Hydrogen rates areexpressed in terms of standard cubic feet per barrel of normally liquidcharge to the reactor, viz. SCFB.

The catalyst used in the hydrotreating reactor should have goodhydrogenating activity. Suitable catalysts comprise a hydrogenatingcomponent as for example the oxide or sulfide of cobalt, nickel, iron,molybdenum, tungsten, chromium, vanadium and mixtures thereof on asupport such as silica, alumina, zirconia,

magnesia and mixtures thereof used as such or in conjunction withzeolites not necessarily of reduced alkali metal content. Preferredcatalysts comprise nickel tungsten on boria-promoted alumina and nickelmolybdenum on activated alumina. The hydrogenating component should bepresent in an amount between about 5 and 40 percent by weight based onthe catalyst composite. Catalysts containing 6 percent nickel and 20percent tungsten or 5 percent nickel and percent molybdenum have beenfound satisfactory.

The temperature within the hydrotreating zone is maintained between 300and 900 F., preferably between 400 and 800 F. Pressure in thehydrotreating zone is substantially the same as that in thehydrocracking zone, taking into consideration the normal pressure droprequired for the flow of materials through the system. Hydrogen shouldbe introduced at a rate of at least 1,000 SCF per barrel of feed, apreferred range being from 3,000 to 15,000 SCFB. The catalyst bed in thehydrogenating reactor should be of a size sufficient to permit liquidhourly space velocities of 0.2-l0 volumes of hydrocarbon liquid pervolume of catalyst per hour. Preferably the LHSV is between 0.5 and 5.

The catalyst used in the hydrocracking stage of our process contains twocomponents, a hydrogenating component supported on a cracking component.The hydrogenating component comprises a Group VIII metal such asplatinum, palladium, iron, cobalt and nickel or compound thereof usedalone or in conjunction with a Group Vl metal such as molybdenum andtungsten or compound thereof. Particularly suitable hydrogenatingcomponents are palladium or nickel and tungsten in sulfide form.

The cracking component of the catalyst comprises a modified crystallinezeolite or a mixture of a modified crystalline zeolite and at least oneamorphous inorganic oxide, the modified zeolite being present in anamount between about 8 and 90 percent by weight. Suitable amorphousinorganic oxides are those displaying cracking activity such as silica,alumina, magnesia, zirconia and beryllia which may have been treatedwith an acidic agent such as hydrofluoric acid to impart crackingactivity thereto. A preferred mixture of amorphous inorganic oxidescomprises silica-alumina in a proportion ranging between 60-90 percentsilica and l0-40 percent alumina.

The modified zeolite portion of the cracking component has uniform poreopenings of from 6-15 Angstrom units, has a silica-alumina ratio of atleast 2.5, e.g. 3-l0, and has a reduced alkali metal content. Themodified zeolite may be prepared by subjecting synthetic zeolite Y toion exchange by contacting the zeolite several times with freshsolutions of an ammonium compound at temperatures ranging between aboutl00 and 250 F. until it appears that the ion exchange is substantiallycomplete. The ion exchanged zeolite is then washed to remove solubilizedalkali metal and dried at a temperature sufficiently high to drive offammonia. This converts the zeolite Y to the hydrogen form and reducesthe alkali metal content to about 2-4 weight per cent. The ion exchangedzeolite is then calcined at a temperature of about l,000 F. for severalhours. After cooling, the ion-exchanged calcined zeolite is subjected toadditional ion exchange by contact several times with fresh solutions ofan ammonium compound and again washed and dried. This treatment resultsin a further reduction of the alkali metal content of the zeolite toless than 1 percent, usually to about 0.5 percent or less. It wouldappear that after the first calcination, it is possible to engage infurther ion exchange with the removal of additional alkali-metal ionsnot removable in the initial ion exchange. Calcination at e.g.1,000l,500 F. may take place here or it may be postponed until after theincorporation of the inorganic oxide and impregnation with thehydrogenating component, at which time the composite should be calcined.Whether the calcination is postponed or repeated, the final calcinationtemperature should not exceed l,200 F.

Hydrocracking catalysts containing a hydrogenating component supportedon a cracking component composed of at least one amorphous inorganicoxide and the twice ion exchanged, twice calcined zeolite have superiorhydrocracking activity and additionally are more resistant todeactivation when brought into contact with nitrogen compounds andpolycyclic aromatics. They also show good stability to steam. Thehydrocracking catalyst should also be substantially free from rare earthmetals and should have a rare earth metal content below 0.5 weight percent, preferably below 0.2 weight per cent. It has been found thatalthough rare earth metals are reputed to enhance the activity andstability characteristics of cracking catalysts, their presence in ahydrocracking catalyst is undesirable.

When the hydrogenating component of the hydrocracking catalyst is anoble metal, it should be present in an amount between 0.2 and 5.0percent by weight based on the total catalyst composite. Preferably thenoble metal is present in an amount between 0.5 and 2 percent. When thehydrogenating component comprises nickel in conjunction with tungsten,the nickel should be present in an amount between about 2 and 10 percentand the tungsten present in an amount between about 5 and 30 percent.Particularly suitable catalysts are those containing between 0.5 and 1.0weight percent noble metal and those containing between 5 and 10 percentnickel and between 15 and 30 percent tungsten. Specific examples ofsuitable catalysts are those containing 0.75 weight per cent palladiumor containing about 6 percent nickel and 20 percent tungsten on asupport made up of about 22 percent modified zeolite Y, 58 percentsilica and 20 percent alumina.

The hydrogenating component is deposited on the cracking component byimpregnating the latter with a solution of a compound of thehydrogenating component. Such techniques are well known in the art andrequire no description here.

1n the hydrocracking reactor, the temperature is generally maintainedbetween about 500 and 900 F., the pressure between 200 and 10,000 psig,the liquid hourly space velocity between 0.2 and volumes of oil pervolume of catalyst per hour, and the hydrogen rate between 1,000 and50,000 SCFB. A preferred temperature range is 600-800 F. Advantageously,the temperature will to a small degree vary depending on the nitrogencontent of the charge, the greater the nitrogen content, the higher thereaction temperature. The preferred pressure range is SOD-3,000 psig.Other preferred conditions are a space velocity of 0.5-2 v/v/hr. and ahydrogen rate of 3,00015,000 SCFB.

The following example is given for illustrative purposes only.

The charge in this example is a light cycle gas oil having the followingcharacteristics:

TABLE 1 Gravity, APl 31.3 ASTM dist. F.

lBP 5% 324-371 l0-20% 429-451 30-40% 479-495 50% 509 60-70% 526-54180-90% 558-580 95-EP 594-614 Total nitrogen, ppm. 63 Basic nitrogen,ppm. 28 Sulfur, wt. 0.29 Polycyclic aromatics, wt. 21.1 Total aromatics,wt. 42.6

The first stage catalyst is composed of 2.2 wt. percent nickel and 10.0wt. percent molybdenum in the form of the oxides and the balance isalumina. The second stage catalyst contains 5.8 wt. percent nickel and19.4 wt. percent tungsten also in the form of the oxides, supported on abase composed of 22 percent decationized zeolite Y, 21 percent aluminaand 57 percent silica. Prior to start-up, the catalysts in both reactorsare sulfided by charging a light petroleum fraction, containing added CSsufficient to bring its sulfur content to 1.0 wt. percent sulfur, at atemperature of 400 F. until the catalysts are sulfided. Sulfiding isthen completed at 600 F. The end-of-run temperatures for the first andsecond stage catalyst are determined to be 750 and 720 F. respectively,a temperature differential of 30 F. The on-stream period is commenced byintroducing the charge in the presence of 5,000 SCFB hydrogen at a spacevelocity (volumes of oil per volume of catalyst per hour) of 2 into thefirst reactor and then passing the entire effluent from the first intothe second reactor where the space velocity is 1.16. Outlet pressurefrom the second reactor is 1,300 psig. At a second reactor temperatureof 652 F. the desired complete conversion to 525 F. and lighter materialis obtained. The temperature in each reactor is gradually raised whileholding the predetermined temperature differential to maintain completeconversion. The product quality is tabulated below:

TABLE 2 Jet Fuel (325-525F.)

Gravity, APl 42.0

Aromatics, vol. X: 14.7

Smoke point, mm. 23 ll5-23SF. naphtha Aromatics, vol. 1.9

RON clear 78.5

RON 3 cc. TEL 95.5 235-325F. naphtha Aromatics 6.9

RON clear 64.0

RON 3 cc. TEL 85.7 Product Yields, fresh feed C C vol. 17.5

C,235 F. vol. 20.3

235-325F. vol. 22.4 Jet Fuel, vol. 56.9 Hydrogen consumption, SCFB 1740By maintaining the designated temperature differential between thecatalyst beds the predetermined end-of-run temperature is reached byboth units at the same time and the entire process is shut down for thesimultaneous regeneration of both catalysts. Each catalyst may beregenerated separately or both catalysts may be regenerated by passingthe regenerating gases serially through the catalyst beds.

The temperature differential is determined to a large extent by theamount of poison present in the feed to the first zone. When thenitrogen content is high, it may be more desirable to increase theend-of-run temperature of the first stage catalyst so that thetemperature differential may be greater than in the case when the feedto the first unit has a lower nitrogen content. For example, if the feedto the first unit is relatively high in nitrogen, as for example,200-2,000 ppm, then the temperature differential may be between 50 andF. whereas if the nitrogen content of the hydrotreater feed isrelatively low, for example, less than ppm, then the temperaturedifferential may be relatively low or possibly even non-existent, forexample 0-30 F. Generally, the end-of-run temperature of thehydrotreating catalyst and the hydrocracking catalyst will range between700 and 800 F. It is also possible in some instances to operate thefirst stage at a lower temperature than the second stage.

Since both the hydrotreating and the hydrocracking reactions areover-all exothermic, there is a temperature rise across each catalystbed. For example, the temperature rise across the hydrotreater catalystbed may be as high as 70-200 F. whereas, with the hydrocracking, thetemperature rise across the individual catalyst beds is generallylimited to 1050 F. in order to maintain good temperature control. Forthis reason, the catalyst bed temperatures mentioned above are theaverage bed temperatures. Since the inlet temperatures are lower thanthe outlet temperatures, the outlet temperature from the hydrotreatermay be 50-100 F. higher than the inlet temperature to the hydrocrackingunit and the first reactor effluent must be cooled prior to itsintroduction into the second reactor. Advantageously, this is done byindirect heat exchange with the fresh feed.

lclaim:

1. In a petroleum hydrocarbon conversion process which is conductedusing two catalyst beds in series in which the first catalyst is apretreating catalyst for controlling the amount of catalyst deactivantentering the second bed and the second catalyst is a conversioncatalyst, the improved method of operation which comprises maintainingsubstantially constant a predetermined average catalyst bed temperaturedifferential between the first and second catalyst beds and controllingthe catalyst bed temperatures to maintain the desired rate of conversionin the second stage.

2. The process of claim 1 in which the first catalyst is a hydrotreatingcatalyst and the second catalyst is a hydrocracking catalyst.

3. The process of claim 1 in which the first and second reaction zonesare maintained at substantially the same pressure.

4. The process of claim 1 in which the second catalyst comprises a noblemetal.

5. The process of claim 1 in which both the first and second catalystscomprise a group Vl metal or compound thereof and a group Vlll metal orcompound thereof.

6. The process of claim 1 in which the petroleum hydrocarbon feed is agas oil and the second stage effects a conversion to jet fuel andlighter material of 50-90 volume percent.

7. The process of claim 1 in which the temperature differential isbetween 10 and 60 F.

8. The process of claim 2 in which the hydrotreating zone is maintainedat a higher temperature than the hydrocracking zone.

9. The process of claim 1 in which the first catalyst comprises a memberof the group consisting of cobalt, nickel and compounds thereof and amember of the group consisting of molybdenum and tungsten and compoundsthereof.

10. The process of claim 1 in which the second catalyst comprises amember of the group consisting of nickel and palladium and compoundsthereof.

11. The process of claim 1 in which both the first and second catalystsare regenerated simultaneously.

12. The process of claim 11 in which each catalyst is regeneratedseparately.

13. The process of claim 11 in which the regeneration is effected bypassing the regeneration gases serially through the catalyst beds.

2. The process of claim 1 in which the first catalyst is a hydrotreatingcatalyst and the second catalyst is a hydrocracking catalyst.
 3. Theprocess of claim 1 in which the first and second reaction zones aremaintained at substantially the same pressure.
 4. The process of claim 1in which the second catalyst comprises a noble metal.
 5. The process ofclaim 1 in which both the first and second catalysts comprise a group VImetal or compound thereof and a group VIII metal or compound thereof. 6.The process of claim 1 in which the petroleum hydrocarbon feed is a gasoil and the second stage effects a conversion to jet fuel and lightermaterial of 50-90 volume percent.
 7. The process of claim 1 in which thetemperature differential is between 10* and 60* F.
 8. The process ofclaim 2 in which the hydrotreating zone is maintained at a highertemperature than the hydrocracking zone.
 9. The process of claim 1 inwhich the first catalyst comprises a member of the group consisting ofcobalt, nickel and compounds thereof and a member of the groupconsisting of molybdenum and tungsten and compounds thereof.
 10. Theprocess of claim 1 in which the second catalyst comprises a member ofthe group consisting of nickel and palladium and compounds thereof. 11.The process of claim 1 in which both the first and second catalysts areregenerated simultaneously.
 12. The process of claim 11 in which eachcatalyst is regenerated separately.
 13. The process of claim 11 in whichthe regeneration is effected by passing the regeneration gases seriallythrough the catalyst beds.